Isothermal-liquid-liquid permeation separation systems



Feb. 20, 1968 c CARPENTER ETAL 3,370,102

ISOTHERMAL-LIQUID-LIQUID PERMEATION SEPARATION SYSTEMS Original FiledFeb. 3, 1966 PERMEATE RICH STREAM ,22 STAGE (N+2) 4 -1 2 go A CONDENSER74 QX EQ STAGE (N+|) GS STREAM A8 EE AR coLD 76 pgg fi ri IOV STREAMWATER REGULATOR 1 STAGE N 2 Y TER FEED 5 R 50 1 STREAM F l G. I ST 5AMPHASE STAGE(N+2) HEAT SEPARATOR ACCUMULATOR 7 EXCHANGER o 48 f 80FEED'TO STEAM LIQUID STAGE (NH) LEvEL OR I CONTROL PRODUCT 56 62 46 \diM FiG. 2 P MAKE UP 160 44 LIQUID RE ULATOR STAGE N STAGE (N+l) I8 EE LL/42 PERMEATE LIQUID PERMEATE 7 LEAN AND T STREAM (5 38 SWEEPIIZIQUID '4036-P as STAGE N '05 Y LIQUID FEED PERMEATE I2 20 LEAN STREAM STAGE L}N+l 34 FEED PREHEATER A STAGE N 32/ 3O STRE R INVENTORS. CLIFFO D LEROYCARPENIER S E RICHARD P. DE FILIPPI KM FGMQ;

' I F l G. 3 ATTORNEY 3,37,102 Patented F eb. 20, 1968 ice 3,370,102HSOTHERMAL-LIQUlD-LIQUID PERMEATION SEPARATION SYSTEMS Clifford LeroyCarpenter and Richard P. de Filippi,

Wellesley, Mass., assignors, by mesne assignments, to Abcor, inn,Cambridge, Mass, a corporation of Massachusetts Continuation ofapplication Ser. No. 524,778, Feb. 3, 1966. This application May 5,1967, Ser. No. 636,546 15 Claims. (Cl. 260674) This application is acontinuation of Ser. No. 524,778 filed Feb. 3, 1966, now abandoned.

Our invention relates to an isothermal process for the separation ofliquids employing membrane permeation techniques. In particular, ourinvention relates to isothermal-liquid-liquid permeation systems whereinliquid mixtures such as hydrocarbons are separated by permeation througha nonporous membrane with the permeate fraction dissolved in a liquidsweep stream and the permeate fraction recovered from the sweep streamby addition of only the latent heat of vaporization of the permeate tothe liquid stream.

The separation of a wide variety of liquid mixtures has beenaccomplished in the past by various membrane permeation techniques. Forexample, hydrocarbon mixtures may be separated into various fractionssuch as aliphatic, aromatic, unsaturated, saturated, straight chain,branch chain and the like, or separated by molecular configuration orboiling points by permeating a portion of the liquid hydrocarbon mixturethrough a nonporous membrane. Separation of mixtures is accomplished bytaking advantage of the difference in the rate at which variousfractions or components of the mixtures permeate a thin, solid,nonporous membrane. In typical processes devised to carry out such aseparation, a feed liquid mixture is placed in contact in a feed zonewith the membrane wherein one fraction of the liquid mixture dissolvesinto the upstream face of the membrane, diffuses through the membranedriven by a concentration gradient and then evaporates from thedownstream face into a permeate zone. The evaporation may be effected bymaintaining a low total pressure downstream from the membrane in thepermeate zone. It is advantageous to remove the permeate fraction fromthe permeate zone rapidly to maintain a good concentration gradient as adriving force. Thus, in some cases a sweep gas is used in the permeatezone at a flow rate sufliciently high to maintain a low partial pressureof the diffusing component, i.e. the permeate fraction. The fraction isthen recovered by condensing the effluent stream from the permeate zone,and

distilling off the sweep gas.

Past liquid-vapor permeation separation techniques have suffered from atleast two major disadvantages resulting from the evaporation of thepermeate in the permeate zone as it passes from the downstream membranesurface. First, the latent heat of vaporization of the permeate fractionmust be supplied at or near the membrane surface. This requires heatingthe feed mixture to the membrane unit, thereby supplying the latent heatfrom the sensible heat of the liquid feed stream. Other means includeintroducing a heat source such as by internally heating the membraneunit by steam or by using a heated sweep gas in the permeate zone. Ineach case, considerable difliculty is and can be experienced ineliminating temperature gradients within the membrane unit, togetherwith the consequent possibility of overheating the thin, nonporoususually polymeric membrane structure. In many cases it is desirable tooperate a membrane system at a temperature approaching or close to themembrane melting point or the critical solution temperature of the membrane, i.e. the melting point of the membrane in contact with anyparticular liquid material. Therefore, providing the latent heat ofvaporization in the permeate Zone or to the membrane surface may resultin damage or destruction of the membrane, a short use life of themembrane material or process operation at considerably less than optimumtemperature conditions.

Another difiiculty associated with the pervaporative systems which havetended to restrict commercial development of membrane techniques is theusually high equipment, power requirements and operating costs necessaryto maintain a suitable vacuum with the total pressure reduced, or forthe compensation of frictional pressure losses in circulating largequantities of a sweep gas at a high flow rate. These and otherdifliculties are particularly evident for a liquid feed mixture of lowvapor pressure at the membrane unit operating temperatures. Theequipment and power cost for vapor or gas compression in such systemscan be expensive and may often be economically prohibitive.

It is an object of our invention to provide a system and means for theseparation of liquid mixtures by an isothermal-liquid-liquid membranepermeation operation.

Another object of our invention is to provide a practical means of andmethod for separating liquid mixtures such as hydrocarbon mixturesthrough the use of a re-cycled liquid sweep stream in an essentiallyisothermal membrane permeation system.

A further object of our invention is to provide a method for theseparation or concentration of various isomers and fractions from liquidmixtures such as the separation of para and meta xylene fractions, andthe separation of butene-l and isobutene fractions from feed streamscontaiaing these mixtures.

Other objects and advantages of our invention will be apparent to thosepersons skilled in the art from the following more detailed descriptionof our invention taken in conjunction with the accompanying drawingwherein:

FIG. 1 is a diagrammatic representation of a multistage membranepermeation system;

FIG. 2 is a schematic illustration of a single-stage membrane permeationsystem for the isothermal separation of a liquid hydrocarbon mixtureemploying a liquid sweep stream; and

FIG. 3 is a schematic illustration of a multi-staged column-typepermeation membrane unit employing hollow fiber-containing membraneplates alternating with solid barriers for each stage.

Briefly, our invention avoids the difficulties associated with thepervaporative and other prior membrane separation systems by theoperation of a liquid-liquid membrane system under essentiallyisothermal conditions and in employing a liquid sweep or solvent streamin the permeate recovery system. The operation of our process at aparticular selected temperature, in combination with the use of a liquidsolvent stream, eliminates or substantially reduces the problem. oftransferring heat to the membrane in the membrane unit, and avoids thehigh equipment and operating costs associated with compression of gasstreams. The use of a liquid sweep stream also is most advantageous inthat in liquid flow only very low pumping energy requirements arerequired to make up for frictional losses, in comparison to that neededin sweep gas flow. Further, to avoid relatively large inputs and outputsof sensible heat, for example through the liquid solvent stream betweenthe membrane unit and other components of the permeate recovery system,the entire system and process is operated in a substantially isothermalmanner.

In operation, a liquid feed stream to be separated is preheated to thedesired isothermal operating temperature of the system and is introducedinto the first stage of a membrane permeation unit divided into a feedzone and a permeate zone. The zones are separated by one or moreselected thin, nonporous membranes. A portion of the liquid feed streamto be removed and recovered diffuses through the membrane, and into thepermeate zone as a fraction enriched in permeate, While the remainingpart of the liquid feed stream, reduced in permeate content, is removedfrom the feed zone and recovered or sent to another or lower stage ofthe membrane permeation system, such lower stage being one with feedstream of reduced. permeate content. A liquid solvent stream, atessentially the same temperature as the feed stream and membrane unitand at a selected flow rate and pressure, is introduced into thepermeate zone, usually but not necessarily countercurrent orcross-current to the flow of the liquid feed stream in the feed zone.The diffused permeate-rich fraction in the permeate zone is dissolved inthe liquid solvent sweep stream.

In a multistage system the liquid feed stream may be mixed with a liquidre-cycled stream from one or more stages or other feed streams ofsimilar composition, such as the permeatelean stream of the next higherstage, and the liquid stream increased in pressure to the operatingpressure of the particular stage to which it is being introduced. Aliquid effluent stream comprising the liquid solvent sweep stream andthe dissolved permeate-enriched fraction is withdrawn from thepermeation zone. This stream is then reduced in pressure to essentiallythe vapor pressure of the permeate-enriched fraction or lower vaporpressure, such as by passing the efiluent stream through a pressurereducing valve controlled by a backpressure regulator. This effiuentliquid stream is then introduced into a heat exchanger such as anevaporator where the latent heat of vaporization for the permeate issupplied to vaporize the permeate-enriched fraction. The vaporizedpermeate-enriched fraction and the still liquid solvent stream, atessentially the isothermal temperature of introduction into the membranepermeation unit, is then introduced into a gas-liquid phase separatorsuch as a cyclone separator. The liquid solvent stream is withdrawn fromthe phase separator. A portion of the sweep stream may be re-cycled backinto the evaporator so that the evaporator may operate with a highpercentage of liquid in the stream without diluting the liquid streameflluent in the permeation zone of the membrane unit. All or theremaining portion of the solvent stream from the phase separator is thenpumped back or increased in pressure to the membrane unit systempressure and reintroduced as the liquid solvent stream to the permeatezone of the membrane unit. Make-up for any liquid solvent lost in thesystem is continuously or periodically added to the recycled solventstream. Sufiicient heat to maintain the selected isothermal temperatureof the process as desired may be added anywhere in the permeate recoverysystem.

Vaporized permeate-enriched fraction withdrawn from the phase separatormay then be recovered by any suitable means. For example, the fractionis condensed and recovered for use as product or sent to an accumulatorvessel in a multistage system, or it may be mixed directly with amixture effluent from a higher stage of the system. Means are providedto maintain the pressure in the recovery system, i.e. the accumulator,the condenser, the phase separator and the evaporator at substantiallythe pressure or slightly lower of the permeate-enriched fracpressure isconsiderably less than the cost of heating and cooling a solvent liquidstream. Our process may be operated as a single stage or as a part of amultistage permeation system, either continuously or as a batch process.The heating of the feed stream is usually employed only in the firststage to bring the liquid feed stream up to the desired temperature,while heat losses in the system can be compensated for anywhere withinthe system so as to maintain the entire system at the optimum orselected constant process temperature. In our system the evaporation ofthe permeate, therefore, takes place due to the difference in pressurein the system and not to any meaningful or deliberate substantialdifferences in temperature within the system.

The selection of the particular isothermal temperature or temperaturerange at which the process is to operate depends in part upon thesolution temperature of the membrane in the presence of permeatingliquid, and the vapor pressure of the permeate-enriched liquid. Ofcourse,

the maximum operating temperature is that at which the.

membrane loses its permeation-selective character at permeationconditions due to dissolution. Below this solution temperature, membranepermeability increases as temperature increases, and consequentlymembrane surface area requirement decreases. However, the vapor pressureof the permeate-enriched liquid and, therefore, the evaporator, phaseseparator, condenser, accumulator operating pressure, also increase withtemperature. Thus,

a reasonable balance in both fixed and operating costs must bemaintained between the investment required for a high-pressure system,versus the amount of membrane surface area required. The selectedisothermal temperature of the system may be about 100 F, for theseparation of high vapor-pressure mixtures, such as of butene-l fromisobutene, or 200350 F., preferably 250 F., for the separation of metaand para xylenes, or higher, depending upon the liquid mixture and themembrane composition. The isothermal temperature of our process mayrange from about 40 to 600, for example to 350 F. The isothermaltemperature selected should be below the critical solution temperatureof the particular membrane employed in the membrane permeation unit,i.e., the melting point of the membrane in contact with the liquid feedstream. The characteristics of the membrane may, therefore, limit theselection of the best isothermal temperature for the system. In othercases, the isothermal temperature selected may be less than the criticalsolution temperature of the system due to the high vapor pressuresinvolved at the higher temperatures, so that economic factors dictatethe use of a lower isothermal temperature. In addition, thepermeation-selectivity of a particular membrane system may vary with thetemperature of the system, so that this factor should be considered inselecting a particular isothermal temperature at which to operate.

The liquid solvent or sweep stream used may comprise any liquid streamhaving certain necessary characteristics. Liquid hydrocarbon streams,due to their ready availability and low cost, are often preferredparticularly when a hydrocarbon mixture is to be separated. A liquidsolvent stream, therefore, may comprise liquid hydrocarbon streams suchas aliphatic, aromatic, unsaturated, saturated streams, as well asvarious naphtha or petroleum fractions, liquefied petroleum gas and thelike. Other liquid solvents which may be employed include but are notlimited to glycols and polyglycols such as ethylene and propyleneglycols and polyols, halogenated hydrocarbons, such asperchloroethylene, fluorocarbons, low molecular weight polymeric liquidsand oils, silicones, water, as well as oxygenated hydrocarbons, such asketones, esters, alcohols and other liquids usually employed inpetroleum or chemical processes as solvents and extraction liquids. Theliquid material selected for the liquid solvent stream should not undulyeffect or degrade the particular membrane system employed, and its backdiffusion through the membrane should be low or negligible. In addition,the permeate-enriched fraction, to be removed or recovered in thepermeate zone, must exhibit good or preferably a high solubility in theliquid solvent stream selected. The liquid solvent stream should have adifferent, such as a lower vapor pressure in comparison to the permeate,i.e. a different or higher boiling point material is. preferred forcases of separation of the permeate enriched fraction. Typically theliquid solvent stream should have a boiling point of at least 50 andoften about 100 or 200 F. or greater than the permeateenriched fractionto be recovered from the permeate zone at the pressure at which permeateis recovered from the solvent. The liquid solvent selected should notnormally form azeotropes with the permeate-enriched fraction.

The permeate-enriched fraction should be recoverable from the solvent byvaporization and preferably by simple, inexpensive flash evaporation ordistillation techniques. In a hydrocarbon separation process, a typicalliquid solvent stream may comprise a kerosene or middle distillatepetroleum fraction having a boiling point range of about 400 to 650 F.,for example a middle distillate of about 500 to 600 F. boiling pointrange.

In the most common useage the liquid solvent or sweep stream should havea lower vapor pressure than the permeate enriched fraction, so that thepermeate-enriched fraction may be easily vaporized and recovered. Insuch an operation the solvent stream is maintained and recycled as aliquid stream, while only the latent heat of vaporization of thepermeate-enriched fraction is furnished to the effluent stream in theevaporator.

- However, the liquid stream used may be more volatile as in those caseswhere the liquid sweep stream represents a small part of the liquideffluent stream withdrawn from the membrane unit. It may then bedesirable to vaporize the sweep liquid. The heat in the evaporator wouldthen be the latent heat of vaporization of the liquid stream and aliquid stream of higher volatility than the permeateenriched fraction isthen chosen. The isothermal conditions of the process are stillmaintained, but the permeateenriched fraction is directly recovered as aliquid stream, while a vaporized sweep liquid is withdrawn from theevaporator and may be recovered such as by condensation and recycled forfurther use in the recovery system. In such cases the efiiuent stream isthen reduced in pressure to the pressure or below of the sweep liquidprior to introduction of the efliuent stream into the evaporator. Atypical sweep liquid would include a liquefied petroleum gas.

The criterion for the selection of a suitable liquid solvent for ourprocess includes low membrane permeability relative to the permeabilityof the mixture to be separated, a different volatility relative to thepermeateenriched fraction, and at least moderate solvent power with thepermeate-enriched fraction. It is also desirable to use a liquid ofreasonably low viscosity to help reduce pumping cost and requirements.

The flow rate of the liquid solvent stream, i.e. the rate of circulationthrough the permeate zone in the membrane permeation unit, is governedin part by the concentration gradient of the permeating liquids which itis desirable to maintain across the membrane. The downstreamconcentration of a permeating liquid can rise to a moderately high levelbefore having a significant effect on the permeation rate. However, evenif very high liquid solvent flow rates are required in our process theeffect on cost of the operating system is low since the major effectwould be on pumping costs which are quite small. This latter effectwould not be true in any nonisothermal process where the heating andcooling requirements would increase in proportion to the increase in theflow or circulation rate. In general, the weight ratio of the liquidsolvent to the permeate in the effluent stream withdrawn from themembrane unit may range from about 0.01 to e.g., 0.1 to 10.0.

The pressure of the liquid solvent sweep stream introduced into thepermeate zone of the membrane unit should be above the vapor pressure ofthe permeate-enriched fraction in the permeate zone and the pressurerequired to pass the stream through the permeate zone. The pressure ofthe liquid stream may range from 10 to 1000 p.s.i.a. Typically thepressure may be the same or a substantially similar pressure to thepressure of the feed stream in the feed zone, particularly whereunsupported, thin membranes are employed. In operation the pressure ofthe liquid efiiuent stream from the permeate zone is reduced prior to oron introduction of the effluent stream into the low pressure recoverysystem, i.e. from the evaporator to the pump in the recycle conduitdiverting the liquid solvent back to the permeate zone. The effluentstream is reduced in pressure to about the pressure or below thepressure of the liquid in the permeate zone which is to be vaporized inthe recovery system. For example, the pressure drop may be up to about200 p.s.i. or more eg. 5100 p.s.i. for a hydrocarbon permeation system.In one preferred operation the pressure of the effiuent stream isreduced just prior to introduction of the stream into a single-stageevaporator to the pressure or just below the pressure of thepermeate-enriched fraction in the permeate zone. Flash evaporation ofthe permeate-enriched fraction then occurs in the evaporator. Largepressure reduction should be avoided to prevent temperature changes inthe isothermal process.

The membrane in the membrane unit may be prepared from any organic orinorganic materials which exhibit selective permeation properties towardthe components of the mixture to be separated. A wide variety ofmaterials are known which exhibit such characteristics. These materialsoften comprise common, natural or synthetic polymeric material which isoften used in a thickness of from about 0.1 to 10 or more mils inthickness, e.g., 1 to 5 mils and may be unsupported or supported. Themembrane material may be used in any sheet, film, tube, hollow fiber orother form which provides a membrane unit having a feed zone and apermeate zone. The membrane material can be used as produced, alone orin combination with other membrane materials, or treated by radiation,solvents, chemicals, orientation or other techniques to enhance theselectivity for the particular separation and/ or the permeation flux.Typical treating methods would include irradiation, chemically reactingthe polymer to change its chemical composition and nature, subjecting itto a solvent swelling and/or thermal cycle, orienting it in a particularfashion or direction by thermal or mechanical stress treatment or byradiation or other means.

Typical membrane materials which may be used include but are not limitedto: C -C polyolefins such as polyethylene and polypropylene; polyamidessuch as nylon; polyesters such as Mylar; fiuoro polymers such as Teflon;acrylic resins; styrene resins such as polystyrene; rubbers such asneoprene, chloroprene, butyl rubber, polybutadiene, copolymers ofbutadiene with styrene, butadiene-nitrile copolymers and other naturaland synthetic elastomers, and cellulose derivatives such as celluloseethers and esters such as hydroxyl cellulose, ethyl cellulose, celluloseacetate and cellulose acetate butarate, vinyl chloride resins, such asSaran, polyvinyl chloride, vinyl chloride-vinyl acetate copolymers,vinyl acetate resins like polyvinyl acetate, silicone rubbers,urethanes, ion exchange resins, glass, ceramics, metal foils and thelike.

The liquid feed material may be any feed mixtures, aqueous ornon-aqueous, separable by semi-permeable membrane techniques. The feedmixtures may be continuously or intermittently introduced into the feedzone. The liquid mixtures may include various petroleum fractions,naphthas, oils, hydrocarbon mixtures, as well as other liquid mixturesincluding chemical reaction mixtures, mixtures of branched andstraight-chain compounds, mixtures of structural, positional and otherisomers, azeotropic mixtures, and the like. For example, our proc essmay be used for the improvement in octane number of gasoline blendingstocks by the selective removal of low octane components from naphthas.Our process is also applicable to the removal of aromatics from keroseneto enhance the smoke point or extraction of normal hydrocarbons to lowerthe freezing point of various petroleum fractions. For example, jet fuelyield may be increased by bringing within the specifications gasolinederived from paraflinic crudes by increasing jet fuel end points, whilemaintaining the necessary product quality. Our process may also be usedfor separating or removing reaction products from a reaction mixture inorder to enhance the selectivity of the reaction or to shift thechemical equilibrium of the reaction.

For the purposes of illustration only our process will be described withregard to the separation of a mixture of meta and para xylenes. In atypical petrochemical process, the product from naphtha reformer isintroduced into a solvent extraction or aromatic recovery unit and a Caromatic fraction removed which comprises a mixture of ortho, meta andpara xylenes and ethylbenzene. This mixture is then distilled to removeethylbenzene and ortho xylene. The meta-para xylene mixture stream isthen introduced into a crystallization unit wherein high-purity paraxylene is recovered, and a mixture of meta and para xylene rejected fromthe unit. This rejected mixture is directed to an isomerization unitwherein an equilibrium mixture of C aromatics is formed and recycledback to the distillation and crystallization units. Our isothermalliquid-liquid membrane permeation system may be employed in such apetroleum process in place of or to supplement the crystallization unit.

FIG. 1 is a schematic illustration of a multistage permeation membranesystem. Three stages, stage N, (N +1) and (N +2) are illustrated here,although more or fewer stages may be employed in any membrane system. Instage N, a feed stream is introduced through conduit 16 into a feed zone12 of a stage N membrane unit 10 with a portion of the feed streamdiifusing as a permeate-enriched fraction into the permeate zone 14 ofthe membrane unit. The permeate-lean fraction stream is withdrawn fromthe feed zone 12 through conduit 20, while the permeate-enriched streamis withdrawn from the permeate zone 14 through conduit'18. Thepermeate-enriched stream may be introduced as all or a portion of thefeed stream into the feed zone of the next higher stage (N +1), which ismembrane unit 20, while the permeatelean stream withdrawn from the feedzone 12 from stage N may be introduced as a part of the feed stream intothe next lower stage (N 1) (not shown). The permeaterich streamwithdrawn from stage (N+l) may be introduced as the feed stream into thenext higher membrane unit 22, stage (N-l-Z).

FIG. 2 is a schematic illustration of a complete isothermalliquid-liquid stage N membrane permeation unit which has beendiagrammatically illustrated in a multistage system in FIG. 1. In oneembodiment of our process a liquid feed stream mixture such as a metaand para xylene mixture, withdrawn from a distillation unit, 1s pumpedthrough conduit lected for the particular membrane system which for thismeta and para xylene separation is approximately 250 F. The heatedliquid stream is Withdrawn through conduit 34 and introduced throughconduit 16 into the feed zone 12 of the stage N membrane unit 10. Thisfeed stream may be mixed with a meta and para xylene stream of similarcomposition, i.e. the permeate-lean stream from stage (N+l) pumped up tothe desired pressure by pump 38 and introduced via conduit 36 andconduit 16 into the feed zone 12. The permeate-lean stream from stage (N+2) may be mixed with the permeate-enriched stream from stage N, and thecombined streams fed to stage (N +1). Optimum operation is obtained whenstage sizes are varied so that only streams of equal or nearly equalcomposition are mixed. The stage N permeation membrane unit 10 isillustrated as divided into a permeate zone 14 and a feed zone 12separated by a semipermeable, thin, nonporous membrane 40, which for theseparation of the meta-para xylene mixture may comprise a 1 milpolypropylene membrane film. The membrane 40 allows selective permeationof para xylene as the liquid feed stream passes from the inlet to theoutlet end of the membrane unit 10.

The para xylene diffuses through the membrane 40 and into the permeatezone 14, while a permeate-lean, i.e..

meta xylene enriched stream is withdrawn from the membrane unit 10through conduit 20. The stage N permeatelean fraction may be introducedas a feed stream into a next lower membrane stage (N-1) or directed backto a crystallization unit.

A liquid sweep stream is introduced through conduit 42 into the membraneunit 10 to pass preferably cross, or as shown, countercurrent in thepermeate zone 14 to the feed stream flow in feed zone 12. The sweepstream is introduced at a pressure about the same as the feed streampressure. The para xylene permeate fraction dissolves in the liquidsolvent stream and the liquid effluent stream with the permeate-enrichedfraction is withdrawn from the membrane unit 10 through conduit 18 whichcontains a pressure valve 44 with a back pressure regulator. This valvereduces the pressure of the efiluent stream to the pressure of thepermeate fraction in the permeate zone or slightly below its pressure.The effluent liquid stream at the reduced pressure is introduced intothe permeate re-.

covery system and into a single-stage evaporator 48. The liquid sweepstream comprises a liquid petroleum fraction having a boiling pointrange of about 500 to 625 F. The meta-para xylene permeate-enrichedfraction is soluble in the solvent stream. The flow rate of the liquidsolvent stream into the permeate zone is controlled so that thecomposition of the effluent stream drawn from the permeate zone 14 asapproximately 50 weight percent solvent and 50 weight percent of thepermeate-enriched fraction. Valve 44 controlled by the pressureregulator lets down the pressure of the liquid solvent eflluent streamfrom a pressure of about 43 p.s.i.a. to the vapor pressure of thepermeate-enriched fraction, i.e., about 9 p.s.i.a. in the.

efiluent stream prior to introducing the liquid stream into the singlestage evaporator 48.

In the evaporator 48 only the latent heat sufiicient to.

vaporize the permeate-enriched fraction is supplied. The vaporizedpern1eateenriched fraction and the liquid solvent stream is then removedthrough conduit 50 and introduced into a phase separator 52 such as agas-liquid cyclone separator. The vaporized permeate-enriched fractionat the isothermal temperature of the system, i.e. about 250 F., iswithdrawn through conduit 64 and introduced into a condenser 66 withcold water coils or 30 to a preheater 32, wherein the stream is heatedto the isothermal temperature se-.

forced air cooling where the permeate-enriched fraction is condensed.The condensed permeate-enriched fraction is then withdrawn from thecondenser 66 through conduit 68 and introduced into an accumulator orstorage vessel 70. A permeate-lean stream from a higher stage, such asstage N +2 and containing the same or a very similar composition as thecondensed permeate-enriched fraction of stage N, may also be introducedvia conduit 72 into the accumulator vessel 70. These condensed fractionsmay be withdrawn from the accumulator through conduit 78 and introducedunder pressure by pump 80 as all or part of the feed stream to the nexthigher stage N+1 or withdrawn as a product stream. Where thepermeateenriched fraction has a low vapor pressure such as in thepresent case with xylenes, a means is employed to maintain the pressurein the recovery system to essentially that of the permeate-enrichedfraction vapor pressure. In the separation of xylenes a system pressureregulation 74 such as a steam ejector communicating through conduit 76to the top of accumulator 70 is employed to maintain the vapor pressurein the accumulator 70, the condenser 66, the phase separator 52 and theevaporator 48 at about the vapor pressure of the permeate-enrichedfraction, i.e. 9 p.s.i.a. Where the permeate-enriched fraction creates apositive gauge pressure in the accumulator 70, the ejector may bereplaced by a pressure regulating valve.

The liquid solvent stream withdrawn from the bottom of the phaseseparator 52 through conduit 54, and still at the isothermaltemperature, is pumped up to the desired operating pressure by pump 58and then reintroduced through the recycle conduit 42 back into thepermeate zone 14. A portion of the liquid solvent stream, for example100 to 1000 percent based on the liquid solvent feed to stage N, may berecycled and introduced into the evaporator 48 through conduit 56 topermit the evaporator to operate with a high percentage of liquidsolvent in the effluent stream without diluting the solventpermeateefiiuent stream in the permeate zone 14. A source of sweep liquid 60 isprovided so that any additional make up liquid required to compensatefor liquid losses in the system, for example by entrainment of theliquid-solvent with the vaporized permeate-enriched fraction going tothe condenser and other losses, may be introduced through conduit 62into the recycle conduit 54. In addition, any heat losses occurringwithin the system can be compensated for at any suitable point in thesystem in order to maintain the isothermal temperature conditions asinitially selected. A liquid level control is employed in conduit 54,which control operates with a control valve in conduit 62, so that thelevel of liquid in the permeate recovery cycle is maintained constant bymake up solvent from source 60.

Our membrane permeation system for the recovery of para and meta xylenemixtures is based on varying the pressure in the permeate recovery cyclerather than the temperature. As is apparent, the cost of energy foroperating, for example, a steam ejector is considerably less than thecost of continually heating and cooling a liquid stream.

In another embodiment of our invention a feed stream comprising a feedmixture containing a C isomeric fraction of butene-l and isobutene maybe separated in our isothermal membrane permeation system as describedemploying a temperature of about 100 F. The isothermal temperature ofthe xylene process was selected based on the upper limit of the criticalsolution temperature of the polypropylene membrane. In the case of thebutylene mixture, a lower temperature is selected in order to keep thecost of equipment at a low level due to the higher vapor pressures ofthe feed stream. In such a system a similar liquid solvent stream andmembrane may be employed as was employed in the xylene operation.

Typical operating conditions for the separation of xylene and butenemixtures by our process in stage N are shown in the following table:

System Butylene Xylene Temperature F.):

a. At membrane b. Permeate Recovery System... Pressure (p.s.i.a.):

a. At membrane b. Penneate Recovery System Feed to Stage N (lbs./hr.):

a. Fresh teed 1). Feed from stage N+l c. Total feed Feed Composition(wt. percent):

a. Butene-l b. Isobutenec. Para xylene d. Meta xylene. Sweep Liquid(lbs/hr.) Permeate-Lean Fraction from Feed Zone- (lbs./hr.)Permeate-Lean Fraction Composition from Feed Zone (wt. percent):

a. Butene-l b. Isobutene c. Para xylene d. Meta xylene"...Permeate-Emiched Fra on 1" m Perme ate Zone (lbs/hr.) Permeate-EnrichedFraction Composition (Swegp tLiquid Free Basis) (wt. percent):

a. u e

FIG. 3 is a schematic representation of a particular multi-stagecolumn-type membrane permeation system. In such a system a column 102contains a plurality of discs acting as a membrane 100, the disc beingseparated by alternating solid barriers 104. Each disc is composed ofvery small diameter hollow fibers of membrane material which give adesirable high surface to volume ratio. In most other cases, the use ofsuch fibers creates pressure drops which may be prohibitively high.However, such high pressure drops are alleviated in part by forming avery squat bundle of low length to diameter ratio, i.e. in the form of adisc or a plate. Such disc or plate to be used, as shown in FIG. 3, isproduced by bundling long, smal1-diameter, hollow fibers together,joining them at closely spaced intervals with a potting resin orcompound, e.g. a liquid which hardens to a solid, and then cuttingthrough the fiber bundle at each point of juncture. Each disc will thencontain a plurality of short fibers connected at common headers. Amultistage system similar to that in FIG. 3 is formed by mounting thesediscs in a column with the fibers essentially vertical. The permeaterecovery system associated with each or any stage has not been shown inFIG. 3. As illustrated, the permeated fluid from the shell side, i.e.the exterior side of the fibers, of stage N is directed to the upstreamheader 105 at the tube side, i.e. the interior side of the fibers, ofstage N-l-l via conduit 106. A nonpermeate or permeate-lean mixture atthe downstream header 107 of stage-N is recycled to the upstream header108 of stage N-l via conduit 109. Similarly, permeate from stage N +1 isfed to the next higher stage (not shown), and the unpermeated portion isrecycled to the upstream header of stage N. An ideal cascade in whichboth streams are of equal or nearly equal composition is produced byprogressively decreasing the thickness of each hollow fiber disc 100 ina calculable fashion as either end of the column is approached.

Of course, any membrane configuration or membrane permeation system maybe used in our process, but a preferred geometry of a membrane would bethat of hollow tubes arranged with common headers in a typicalshell-and-tube type exchanger or vessel. For a larger diameter tube, sayA; inch to 1 inch or more, a simple, straight bundle with or withoutbafiies can be employed. Where smaller diameter tubes are used, say inthe range of of an inch to of an inch, it is often necessary to use aspecial arrangement in order to insure a uniform distribution of flow onthe shell side of the fiber bundle. As in our process the multistagecolumn system may be maintained under isothermal conditions bysurrounding the column and conduits with an appropriate insulation,controlling the temperature of the feed streams and making up any heatlosses occurring during the separation process.

Our process has been described for purposes of illustration onlyemploying xylene and butene type hydrocarbon isometric mixtures.However, our process is applicable to any liquid mixtures, both aqueousand nonaqueous, capable of being separated by a membrane permeationsystem. The use of an isothermal system whereby the temperature of thepermeation unit and the permeate recovery system is maintained constantwith variations only in the pressure of the system provides a system oflow operating cost. In addition, the use of a liquid solvent stream insuch an isothermal system, with a single-stage evaporation unit withrecycle of the liquid sweep stream, avoids many of the difficultiesassociated with the prior art. Our process may be operated as a singlestage membrane process or incorporated into other known petroleum orchemical process operations, or preferably operated as a multistagemembrane permeation process.

What we claim is:

1. A membrane permeation method for the separation of liquid mixtureswhich method comprises:

introducing a liquid feed mixture to be separated into apermeate-enriched and a permeate-lean liquid fraction into a feed zone,the feed zone being part of a membrane permeation system characterizedby a permeate zone and a feed zone, the zones separated by a thinmembrane material exhibiting a selective permeation affinity for atleast one component of the feed mixture;

permeating a fraction of the liquid feed mixture through the membranefrom the feed zone into the permeate zone to obtain in the permeate zonea permeate-enriched fraction; contacting the permeate-enriched fractionin the permeate zone with a liquid sweep stream, which liquid sweepstream is a solvent for the permeate-enriched fraction, and has adifferent volatility than the permeate-enriched fraction and from whichthe resulting liquid efiiuent stream the said fraction and liquid may beseparated by evaporation; withdrawing from the permeate zone a liquidefiluent stream comprising as components thereof the permeate-enrichedfraction and the liquid sweep stream;

withdrawing from the feed zone a permeate-lean fraction;

reducing the pressure of the liquid effluent stream to a pressure ofabout the pressure or lower in the permeate zone of the component of theliquid efiluent stream to be vaporized;

vaporizing one component of the liquid effiuent stream by addition ofonly the latent heat of vaporization for that component;

separating the vaporized component and the liquid component of theeffluent stream;

recovering the permeate-enriched fraction; and

recycling at least a part of the liquid sweep stream to the permeatezone, all steps of the process being carried out under a selected singlesubstantial isothermal condition without intentional steps to change thetemperature of the liquid sweep, feed, or liquid effluent streams.

2. The method of claim 1 which includes:

employing a liquid sweep stream having a lower volatility than thepermeate-enriched fraction in the permeate zone;

reducing the pressure of the liquid effluent stream to about thepressure or less of the permeate-enriched fraction in the permeate zone;

vaporizing substantially only the permeate-enriched fraction bysupplying the latent heat of vaporization to the reduced pressure liquideflluent stream;

condensing the vaporized permeate-enriched fraction;

and

withdrawing the condensed permeate-enriched fraction as a productstream.

3. The method of claim 2 which includes:

maintaining the pressure in the vaporizing, separating,

condensing and withdrawing of the condensed permeate-enriched fractionsteps of the permeate re covery at about the pressure of thepermeate-enriched fraction in the permeate zone.

4.. The method of claim 1 wherein the liquid solvent sweep stream ischaracterized by a boiling point of greater than 50 F. or above theboiling point of the permeateenriched fraction in the permeate zone.

5. The method of claim 1 wherein the feed stream includes apermeate-lean fraction stream withdrawn from another stage of amulti-stage permeation membrane process.

6. The method of claim 1 wherein the liquid feed stream comprises amixture of hydrocarbon stream.

7. The method of claim 1 wherein the membrane includes a plurality ofsmall diameter hollow fibers fabricated into bundles of lowlength-to-diameter ratio, as membranes between a permeate and feed zone.

8. The method of claim 1 wherein the process represents one stage in amulti-stage membrane permeation system and which includes introducingthe permeate-lean fraction withdrawn from the feed zone into the feedzone of a stage lower in content of the more permeable component andintroducing the condensed permeate-enriched fraction into the feed zoneof a stage higher in content of the more permeable component.

9. The method of claim 1 which includes:

introducing the liquid sweep stream into the permeate zone at a pressureof about the pressure of the liquid feed mixture in the feed zone.

10. The method of claim 1 which includes:

employing a liquid sweep stream having a higher volatility than thepermeate-enriched fraction in the permeate zone;

reducing the pressure of the liquid effluent stream to about thepressure of the liquid sweep stream or lower;

vaporizing the liquid sweep stream;

recovering the liquid permeate-enriched fraction; and

condensing the liquid sweep stream.

11. The method of claim 1 which includes:

withdrawing a liquid solvent stream and a vaporized permeate-enrichedfraction stream from a single stage evaporator wherein the latent heatof vaporization of the permeate-enriched fraction is supplied;introducing the stream into a gas-liquid phase separator; withdrawingfrom the phase separator a vaporized permeate-enriched fraction;condensing the vaporized permeate-enriched fraction, and withdrawing theliquid solvent stream from the phase separator;

introducing a portion of the liquid stream to the evaporator; and

recycling a portion of the liquid solvent stream into the permeate zone.

12. The method of claim 11 which includes:

reducing the pressure of the liquid effluent stream to below thepressure of the permeate-enriched fraction in the permeate zone wherebyflashing of the permeate-enriched fraction occurs upon introduction ofthe liquid stream into the evaporator.

13. The method of claim 11 which includes:

introducing the liquid sweep stream at a pressure of between 10 and 1000p.s.i.a. into the permeate zone in a different flow direction than theflow direction of the feed stream, the weight ratio of liquid sweep 1314 stream to the permeate-enriched fraction ranging the feed mixturecomprising a mixture of C hydrofrom 0.1 to 10. carbon isomers, to beseparated.

14. The method of claim 11 which includes:

maintaining the temperature of the process at a con References Citedstant temperature between about 80 and 350 F., the 5 UNITED STATESPATENTS feed mixture comprising a mixture of meta and para xylenes, tobe separated 3,244,763 4/ 1966 Cfl n 7 15. The method of clalm 10 whichincludes: DELBERT E. G ANTZ, Primary Examiner maintaining thetemperature of the process at a constant temperature of between about 40and 200 F 10 C. E. SPRESSER, Assistant Examiner.

1. A MEMBRANE PERMEATION METHOD FOR THE SEPARATION OF LIQUID MIXTURESWHICH METHOD COMPRISES: INTRODUCING A LIQUID FEED MIXTURE TO BESEPARATED INTO A PERMEATE-ENRICHED AND A PERMEATE-LEAN LIQUID FRACTIONINTO A FEED ZONE, THE FEED ZONE BEING PART OF A MEMBRANE PERMEATIONSYSTEM CHARACTERIZED BY A PERMEATE ZONE AND A FEDD ZONE, THE ZONESSEPARATED BY A THIN MEMBRANE MATERIAL EXHIBITING A SELECTIVE PERMEATIONAFFINITY FOR AT LEAST ONE COMPONENT OF THE FEED MIXTURE; PERMEATING AFRACTION OF THE LIQUID FEED MIXTURE THROUGH THE MEMBRANE FROM THE FEEDZONE INTO THE PERMEATE ZONE TO OBTAIN IN THE PERMEATE ZONE APERMEATE-ENRICHED FRACTION; CONTACTING THE PERMEATE-ENRICHED FRACTION INTHE PERMEATE ZONE WITH A LIQUID SWEEP STREAM, WHICH LIQUID SWEEP STREAMIS A SOLVENT FOR THE PERMEATE-ENRICHED FRACTION, AND HAS A DIFFERENTOLTILITY THAN THE PERMEATE-ENRICHED FRACTION AND FROM WHICH THERESULTING LIQUID EFFLUENT STREAM THE SAID FRACTION AND LIQUID MAY BESEPARATED BY EVAPORATION; WITHDRAWING FORM THE PERMEATE ZONE A LIQUIDEFFLUENT STREAM COMPRISING AS COMPONENTS THEREOF THE PERMEATE-ENRICHEDFRACTION AND THE LIQUID SWEEP STREAM; WITHDRAWING FORM THE FEED ZONE APERMEATE-LEAN FRACTION; REDUCING THE PRESSURE OF THE LIQUID EFFLUENTSTREAM TO A PRESSURE OF ABOUT THE PRESSURE OR LOWER IN THE PERMEATE ZONEOF THE COMPONENT OF THE LIQUID EFFLUENT STREAM TO BE VAPORIZED;VAPORIZING ONE COMPONENT OF THE LIQUID EFFLUENT STREAM BY ADDITION OFONLY THE LATENT HEAT OF VAPORIZATION FOR THAT COMPONENT; SEPARATING THEVAPORIZED COMPONENT AND THE LIQUID COMPONENT OF THE EFFLUENT STREAM;RECOVERYING THE PERMEATE-ENRICHED FRACTION; AND RECYCLING AT LEAST APART OF THE LIQUID SWEEP STREAM TO THE PERMEATE ZONE, ALL STEPS OF THEPROCESS BEING CARRIED OUT UNDER A SELECTED SINGLE SUBSTANTIAL ISOTHERMALCONDITION WITHOUT INTENTIAONAL STEPS TO CHANGE OTHE TEMPERATURE OF THELIQUID SWEEP, FEED, OR LIQUID EFFLUENT STREAMS.